Synthesis of organic compounds from carbon monoxide and hydrogen



H. G. MCGRATH ETAL June 19 1956 SYNTHESIS oF ORGANIC COMPOUNDS FROM2,751,405

CARBON MONOXIDE AND HYDROGEN Filed April 9, 1951 awww/"ff Was. amm/ H/sH Tof/vayas United States Patent O SYNTHESIS OF ORGANIC COMPOUNDS FROMCARBON MONOXIDEAND HYDROGEN Henry G. McGrath, Union, and Louis C. Rubin,West Caldwell, N. I., assignors to The M. W. Kellogg Company, JerseyCity, N. J., a corporation of Delaware Application April 9, 1951, SerialNo. 220,050 1 Claim. (Cl. 260--449.6)

This invention relates to an improved method for hydrogenating carbonoxides to produce hydrocarbons and oxygenated organic compounds.Primarily the improved process involves reacting hydrogen and carbonmonoxide under highly eflicient conditions to produce hydrocarbons andoxygenated organic compounds under which conditions the major proportionof the oxygen of the carbon oxide is converted to water rather thancarbon dioxide. The improved process is applicable also in reactinghydrogen with organic compounds containing the carbonyl group, andherein designated generally as carbon oxide reactants, whose reactionwith hydrogen is promoted by the catalysts which are effective withcarbon monoxide, such as ketones, aldehydes, acyl halides, organic acidsand their salts and esters, acid anhydrides, and amines. In thefollowing description of the invention the hydrogenation of carbonmonoxide will be referred to specifically. It will be understood,however, that the invention is of Wider application, including withinits scope the hydrogenation of any suitable carbon oxide. f

This application is a continuation-in-part of my prior and Vcopendingapplications Serial No. 690,820, filed August 15, 1946, and Serial No.729,878, tiled February 20, 1947, both now abandoned.

'Ihe improved process involves flowing a gaseous mixture comprisinghydrogen and the carbon oxide to be hydrogenated upwardly in a reactionzone containing a mass of finely divided metal catalyst for thereaction. The hydrogen and carbon oxide reactant are passed in the gasform through the reaction zone, under conditions effective to react all,or a major proportion, of the carbon oxide reactant, at a charging rate,in relation to the quantity of catalyst in the reaction zone, which ismuch higher than the ycharging rates previously employed in similaroperations. The gaseous mixture is passed upwardly through the mass ofcatalyst at a velocity effective l to suspend or entrain the catalystmass in the gas stream. The velocity of the gas stream passing throughthe reaction zone, however, is preferably suiiiciently low to maintainthe catalyst mass in a dense, uidized, pseudoliquid condition. In thiscondition the catalyst mass may be said to be suspended in the gasstream, but not entrained therein in the sense that there is movement ofthe catalyst mass as such in the direction of flow of the gas stream. Itis preferred, however, to maintain the upward velocity of the gas streamsuiiciently high 2,751,405 Patented June 19, 1956 process is defined byreference to the rate at which the total gaseous feed is charged, interms of standard cubic feet, in the gas form, per hour, per pound ofthe metal catalyst in the reaction-zone. The improved process isoperated at `a minimum space velocity equivalent to chargingat least31,0v standard cubic feet of total gaseous feed, per hour, per pound ofthe metal catalyst in the catalyst phase of the reactor. A standardcubic foot of gas is that quantity -of normally gaseous components whichwould occupy one cubic foot at atmospheric pressure and F., or anequivalent quantity of normally liquid feed components. When reactingcarbon monoxide it is preferred to employ still higher spacevelocities,`as will be described in'more detail below.

The catalyst employed in the present invention is a finely dividedpowder `consisting essentially of an iron, or iron oxide, which is, orbecomes in the reaction zone, a catalyst for the reaction, or a mixtureof such iron and iron'oxide'catalytic materials.- While the catalystpowder consists essentially of such catalytic iron and iron oxides itmayinclude also a minor amount of promoting ingredients, such asalkalies, alumina, silica, titania, thoria, manganese oxide, magnesia,etc. ln the following description and claim, catalyst powders consistingof a metal and/or a metal oxide and containing at most a minorproportion ,of promoter are referred to as finely divided metalcatalyst. f

The exact chemical condition of the catalyst in yits most active form isnot certain. It may be that the active form is present when the metal isin an optimum degree of oxidation or carburization. Consequently, thepowdered catalyst which is in la` substantially completely reducedcondition when'lirst contacted with the reactants, may reachk vits stateof highestv activity through being oxidizedfand/l or carburized in thereaction zone. Therefore, in this specification and claim the catalystemployed is described by` reference to its chemical condition when rstcontacted with the reactants.

The catalyst is employed in a fine state of subdivision. Preferably thevpowdered catalyst initially contains no more than a minor 'proportionby weight of material whose particle size is greater than 250 microns.Preferably also the greater l,proportion of the catalyst mass comprisesmaterial whose particle size is smaller than 100 microns', including yatleast 25 weight per cent of the material in particle sizes lsmaller than37 microns. A highly desirable powdered catalyst comprises at least 75per cent by weight of material smaller than 150 microns in 'particlesize, and at least 25 per cent by weight smaller than 37 microns inparticle size. v

In the preferred form of the invention the powdered catalyst mass ismaintained in a reactor substantially larger than the volume occupied bythe catalyst mass in the iiuidized condition.v 'In this voperation allbut a minor proportion of the catalyst mass is contained in the dense tomaintain the uidized catalyst mass in a highly turbulent condition inwhich the catalyst particles circulate at a high rate in thepseudo-liquid mass. In this preferred condition of operation a smallproportion of catalyst in the iiuidized mass may become entrained in thegas stream emerging from the upper surface of the fluidized mass wherebycatalyst thus entrained is carried away from the mass.

In the improved process the hydrogen and carbon oxide are employed inratios such that there is a substantial excess of hydrogen.` Thecharging rate inthe improved iiuidized pseudo-liquid mass, which may bedesignated as the dense phase of the catalyst. The dense phase of thecatalyst occupies the lower part of the reactor While that part of-thereactor above the dense phase is occupied by aV mixture of gases andpowdered catalyst in which the concentration of catalyst is much lower,vand of a different order of magnitude, than the concentration of thecatalyst in the dense phase. This diffuse phase may be said to beadisengaging zone in which the solids lifted above the dense phase bythe gas stream are disengaged therefrom and returned to the dense phaseto the extent that such solids are present in the diffuse phase inexcess of the carrying capacity of the gas stream at the superi 'cialvelocity f the gas stream. The latter is the velocity at which the gasstream would flow through the reactor in the absence of catalyst. In thedense phase the concentration of the catalyst in the gas stream variesfrom a maximum near the gas inlet to a minimum in the upper part of thisphase. Likewise the concentration of catalyst in the diffuse phasevaries from a maximum near the upper surface of the dense phase to aminimum in the upper part of the reactor. Between the dense phase ofhigh average concentration and the diffuse phase of low averageconcentration there is a relatively narrow zone in which theconcentration of solids in the gas stream changes in a short space `fromthe high concentration of the dense phase to the low concentration ofthe diffuse phase. This zone has the appearance of an interface betweentwo visually distinct phases.

As the improved method of operation ordinarily involves employment ofcatalyst powders and gas velocities such that a portion of the densefluidized catalyst mass is carried away by entrainment, it is necessaryto provide means in the reactor for separating such entrained catalystand returning it to the dense phase, or to provide means externally4 ofthe reactor to separate entrained catalyst from the gas stream andreturn it to the reactor, or otherwise to recover catalyst from theproduct gas stream. When catalyst is permitted to pass out of thereactor in entrainment in the gas stream it is necessary to return suchcatalyst to the reactor, or replace it with fresh or revivifiedcatalyst, in order to maintain the desired volume of fiuidized catalystin the reaction zone.

The improved method of operation, in which the finely powdered catalystis employed in a form consisting of the metal catalyst, or its oxide,and containing at most minor proportions of promoting agents, providesvery high catalyst concentrations in the reaction zone. The employmentof the finely powdered metal catalyst in a fluidized bed with eflicientcooling means also is a factor in 'permitting the use of high catalystconcentrations, since it facilitates the removal of heat from therelatively concentrated reaction zone. The improved operation, employingthe finely divided metal catalyst, results in initial catalystconcentrations of at least 30 pounds per cubic foot of the fluidizeddense catalyst phase, while the preferred gas velocities result ininitial concentrations of 40 to 120, or more, pounds per cubic foot ofdense phase. It will be understood that these figures refer to theinitial average concentration in the dense phase. The accumulation ofreaction products on the catayst particles as the operation proceedsreduces the catalyst density and increases the bulk of the denseiiuidized mass.

The temperature employed may approximate those employed with thecatalyst in question in fixed catalyst bed operations. With the ironcatalyst, temperatures in the range of 450 F. to 750 F. are employed,while average temperatures between 550 F. and 650 F. are preferred.Likewise the pressures employed may approximate those previouslyemployed in fixed bed operations. With the iron catalyst, for example,pressures between atmospheric pressure and the maximum pressure at whichcondensation is avoided may be employed. It is desirable, however, toemploy pressures of at least 150 p. s. i. g., preferably between about200 and about 600 pounds per square inch gage.

In this specification, pressures are expressed as pounds per square inch(gage) and gas volumes as cubic feet measured at 60 F. and atmosphericpressure.

The linear velocity of the gas stream passing upwardly through the densephase is conveniently expressed in terms of the superficial velocity,which is the linear velocity the charge gas stream would assume ifpassed through the reactor in the absence of catalyst. This isdesignated 4 either as inlet velocity or average superficial velocity,the latter taking into account the shrinkage in Volume caused by thereaction. These velocities preferably are in the range of from 0.1 to 10feet per second, but higher velocities may be used without departingfrom the scope of this invention.

In a preferred modification of the invention a metallic iron powder,having the preferred distribution of particle sizes and having combinedtherewith a small amount of promoters, such as alkalies andnon-reducible oxides, is employed under special conditions of operationto effect rates of conversion not previously attained and to convert themajor proportion of the oxygen of the carbon oxide reactant to waterwhereby only a minor proportion of the oxygen of the carbon oxidereactant is converted to carbon dioxide.

The improved process is carried out at space velocities substantiallygreater than those previously employed.y The reactants are passed intoand through the reaction zone at a space velocity equivalent to at least30 standard cubic feet of total gaseous feed, per hour, per pound ofmetal catalyst in the catalyst phase of the reactor. In thehydrogenation of carbon monoxide with an iron catalyst, it is preferredto operate also at a space velocity equivalent substantially above 4,preferably at least 10, standard cubic feet of carbon monoxide, perhour, per pound of iron catalyst in the dense catalyst phase. In theimproved process the mol ratio of hydrogen to carbon monoxide is atleast 2.521 and preferably at least 5:1. The maximum space velocity islimited principally by the capacity of the equipment and the ability forremoving the exothermic heat of reaction from the reaction zone. Inpseudo-liquid dense phase catalyst operation, it is preferred to limitthe maximum space velocity to an amount equivalent to 15 standard cubicfeet of carbon monoxide, per hour, per pound of catalyst in the reactor.

The volume of reactants, per hour, per volume of dense catalyst phasedepends upon the charge rate and also upon the density of the densephase, the latter being affected by the condition of the catalyst andthe gas velocity. At the preferred gas velocities mentioned above, andwhen employing the iron catalyst, the minimum space velocity withrelation to carbon monoxide may be defined as about 400 volumes ofcarbon monoxide (measured at standard conditions of temperature andpressure) per hour per volume of the dense catalyst phase. The volume ofdense catalyst phase is that occupied by the catalyst when uidized in afresh condition by the charge gas at the reaction velocity. The absolutespace velocity of the total charge gas, which is at least 30, affectedby the hydrogen to carbon monoxide ratio as well as by the presence ofother constituents, such as nitrogen, carbon dioxide, and hydrocarbongases. The reaction gas mixture may include, in addition to hydrogen andcarbon oxide reactant, other relatively non-reactive ingredients, suchas nitrogen, and hydrocarbon gases, such as methane, ethane and propane.

The operation is carried out with a charge gas containing hydrogen andcarbon oxide reactant in a ratio substantially greater than the ratio inwhich these compounds are converted to other compounds in the reactionzone. Previous investigators have noted little advantage in the use ofH2:CO ratios greater than 1:1 in connection with iron catalysts. In thisimproved process it has been discovered, however, that materialadvantages follow the use of H2:CO ratios greater than about 23:1,preferably greater than about 5: 1. The presence of excess hydrogen inthe reaction zone favorably atects the quality of the product, improvesthe selectivity ofthe reaction, minimizes the formation of carbon andthus facilitates operation at high temperature levels, lessens theformation of carbon dioxide, and minimizes the need for revivilicationof the catalyst. t A In connection with the present invention it hasbeen discovered that the conversion operation carried out in the mannerdescribed above can be extended substantially indefinitely without thenecessity for catalyst regeneration by careful control of the ratio ofhydrogen to carbon monoxide. In the foregoing operation the metalcatalyst accumulates carbonaceous deposits including tarry material,waxy materials, hydrocarbon liquids and oxygenated compounds of highmolecular Weight. lt is found that these deposits continue to accumulateon the catalyst at a rate and to a final percentage which is affected bythe temperature and the ratio of hydrogen to carbon monoxide in thecharge gas mixture. It has been found, when operating at temperatureseffective for a high conversion rate, that the lower the ratio ofhydrogen to carbon monoxide in the charge gas mixture the more rapidlwill be the accumulation of carbonaceous deposits on the catalystsurface and the higher will be the percentage of the catalyst massrepresented by carbonaceous deposits when equilibrium conditions arereached. More specifically, it has been found that if the mol ratio ofhydrogen to carbon monoxide is maintained greater than :1 theaccumulation of carbonaceous deposits is stabilized at a relatively lowpercentage of the total catalyst mass whereby the activity of thecatalyst under such stabilized conditions of operation is sufficientlyhigh to continue the operation indefinitely without the necessity forrevivification treatment of the catalyst. More specifically, it has beenfound that the operation can be continued indefinitely without catalystregeneration if the H2:C0 ratio is maintained greater than 5:1, forexample, about :1, or greater, but not higher than about 12:1.

Conveniently, the high HzzCO ratio may be maintained in the charge gasmixture by the combined effect of supplying a fresh feed gas mixturecontaining hydrogen and carbon monoxide in a higher mol ratio than theratio at which these components are reacted and recycling unconvertedgases to form a composite charge gas mixture having the desired ratio ofhydrogen to carbon monoxide. In the hydrogenation of carbon monoxide bythe manner described above the carbon monoxide content of the charge gasmixture is substantially completely reacted during the passage of thecharge gas through the reaction zone. 'Consequently the unconverted gascontains hydrogen in an H2:CO ratio substantially greater than in thecharge gas. By recycling such unconverted gas in combination with afresh feed containing H2 and COin a lower ratio than that desired in thecharge gas a cornposite charge gas mixture having the desired-ratio'may5' be prepared.

The gases to be recycled may be obtained from the reaction product by asimple preliminary cooling of the product which separates only the morereadily separable liquid reaction products, or the recycle gases may beobtained from the product gas after extensive condensation treatment toremove substantially all condensable hydrocarbons and oxygenatedcompounds.- yThe vvolumetric ratio of recycle gas to fresh feed gas ispreferably between about 0.5 :l and about 5:1, but other ratios maybeused if desired, for example, ratios as low as 0.3:1 and as high as10:1. A composite charge gas mixture may be formed by combining a freshfeed mixture having an HzzCO' ratiol of about 3:1 with a recycle gasstream containing essentially no carbon monoxide to produce a compositefeed containing hydrogen and Vcarbon monoxide in a ratio of about 12: 1.The carbon monoxide in this composite feed is substantially completelyreacted in passing the mixture through the reaction zone, usually above95 per centconversion on an over-al1 basis. A

.6 portion of the unconverted gases is discarded from th system toprevent the accumulation of inerts, such -as nitrogen, and theremaindermay be: recycled in an amount effective to produce the desiredratio in the composite charge gas.

The principal effect of a high HzrCO ratio on the reac'tion product isthe substantial elimination of CO2 as a product. In previous operationsemploying H2:CO ratios heretofore considered desirable, about 40 percent of the CO converted appeared in the product as'COz. In the practiceof this invention it has been possible to avoid any production of CO2and even effect consumption of CO2 in the feed gas, and a majorproportion of the oxygen of the carbon monoxide is converted to water.

In operating a synthesis process in accordance with the presentinvention in which the normally gaseous products include carbondioxide,`the total inlet feed contains vat least 6 or 7 volume per centcarbon dioxide.v It is preferred, however, to limit the carbon dioxidecontent of the total inlet feed including recycle to less than 25 volumeper cent. Particularly good .results have been observed with the carbondioxide concentration of the total feed inlet gas between about 8 andabout 14 volume per cent. The fresh feed will usually contain less than6 volume per cent carbon dioxide and the carbon dioxide content of thetotal charge gas is thus made up by recycle. After the recycle has builtup the carbon dioxide content of the total feed gas within the abovelimits, the carbon dioxide, under the preferred conditions, will beconsumed at a' rate substantially equivalent to the amount of carbondioxide charged and thus eliminating the net production of carbondioxide and rejection of the combined oxygen of the feed as water ratherthan carbon dioxide. `Some of the oxygen will be converted to oxygenatedorganic compounds but the major proportion of the oxygen is converted towater.

The rejection of oxygen from the system as water rather than carbondioxide is particularly advantageous as it minimizes the size andcomplexity of subsequent treating equipment by eliminating the treatmentof the effiuent to remove carbon dioxide and by minimizing venting of aportion of the product gas. The water is removed from the eluent bysimple condensation at operating pressure or lower.

Suitable contact times between about 8 and about 95 seconds have beenfound satisfactory, preferably between 14 and 35 seconds.

If regeneration or revivification of the catalyst is required,regeneration may be accomplished by mere stripping with hydrogen, carbondioxide, steam or other gas, or by oxidationavyith air, steam or oxygenat atemperav"turebetween about 600 F. and 2000 F. or by reduction withhydrogen at a temperature between about 500 F. and about l000 F., or byany combination of the above. Regeneration may be effected in separateequipment by removal of spent catalyst, or in the same equipment usedfor synthesis by discontinuing the flow of reaction gas and substitutionof regeneration gas.

The invention will be described further by reference to the'accompanying vdrawing which is a view in elevation, partly in section,of a reactor employed in carrying out the present invention, and byreference to specific examples of operations embodying the presentinvention and carried out in apparatus exemplified by the drawing.

Referring to the drawing, reactor 11 consists of a length of extra heavystandard 2-inch steel pipe which isA about 153 inches long and hasinside and outside diameter o-f 1.94 inches and 2.38 inches,respectively. Reactor 11 is connected, by conical section 12, to aninlet pipe 13 made of extra heavy standard half-inch steel pipe havingan inside diameter of 0.55 inches. Reactor 11 is connected at the top,by means of conical section 14, with an enlarged conduit comprising alength of 6-inch extra heavy standard steel pipe having an insidediameter of 5.76 inches. Conical section 14 and conduit 15 constitute anenlarged extension of reactor 11 which facilitates disengagement ofcatalyst from the gas stream after passage of the latter through a densecatalyst phase.

Conduit 15 is connected by means of manifold 16 with conduits 17 and 18which comprise other sections of extra heavy 6-inch standard steel pipe.Conduits 17 and 18 contain filters 19 and 20 which are constructed ofporous ceramic material which is permeable to the gas and vaporsemerging from the reaction zone but impermeable to the catalystparticles carried by entrainment in the gas stream. Filters 19 and 20are cylindrical in shape and closed at the bottom ends. They aredimensioned in relation to conduits 17 and 1S to provide a substantialannular space between the filter and the` inner wall of the enclosingconduit for the passage of gasesfand vapors and entrained catalystupwardly about the outer surface of the filter. The upper ends of lters194 and 20 are mounted in closure means 21 and 22 in a manner wherebythe gases and vapors must pass through either filter 19 or ilter 20 toreach exit pipes 23 and 24. Each of filters 19 and 20 is approximately36 inches long and 41/2 inches in outside diameter, the ceramic filterwalls being approximately 3A of an inch thick.

The greater part of reactor 11 is enclosed in the jacket which extendsfrom a point near the top of the reactor to a point sufficiently low toenclose the 3 inch length of conical section 12 and approximately 5inches of pipe 13. Jacket 25 comprises a length of extra heavy 4-inchstandard steel pipe having an inside diameter of 3.83 inches. The endsof jacket 25 are formed by closing the ends of the 4-inch pipe in anysuitable manner, as shown, and sealed by welding. Access to the interiorof jacket 25 is provided by an opening 26 in the top thereof through a2-inch steel pipe. Jacket 25 is adapted to contain a body of liquid fortemperature control purposes, such as water, or Dowtherm (diphenyl ordiphenyl oxide or a mixture of same). The vapors which are evolved bythe heat of reaction in reactor 11 are withdrawn through conduit 26,condensed by means not shown, and returned through conduit 26 to thebody of temperature control iluid in jacket 25. Electrical heating means(not shown) is provided in connection with jacket 25 to heat thetemperature control fluid therein to any desired temperature, for useparticularly when starting up the hydrogenation reaction;

a In order to show all the essential parts of the reactor and associatedcatalyst separation means on a single sheet a large proportion of theapparatus has been eliminated by the breaks at 27 and 28. For a clearunderstanding of the relative proportions of the apparatus reference maybe had to the over-all length of the apparatus, from the bottom ofjacket 25 to exit pipes 23 and 24, which is about 224 inches. In each ofbreaks 27 and 2S the portion of the apparatus eliminated is identicalwith that portion shown immediately above and below each break.

In the operations carried out in the apparatus of the drawing, thecatalyst recovery means comprising filters 19 and 20 is effective toseparate substantially completely en trained catalyst from the outgoingstream of gases and vapors. The disengagement of solids from the gasstream is promoted by the lowered velocity of the gas stream in conduit15 and remaining solids are separated on the outer surfaces of filters19 and 20. The latter are employed alternately during the operation sothat the stream of gases and vapors and entrained solids passes fromconduit 15 through either the left or right branches of manifold 16 intoeither conduit 17 or conduit 18. During 8 the alternate periods thefilter which is n ot in use is subjected to a back pressure of gas whichis introduced at a rate sufficient to dislodge catalyst which hasaccumulated on the outer surface of the filter during the active period.Such blowback gas and dislodged catalyst flow downwardly in the conduitenclosing the filter and into manifold 16 in which the blowback gas iscombined with the reaction mixture owing upwardly from conduit 1S. Thegreater part of the catalyst thus dislodged settles downwardly into thereactor and is thus returned for further use. The blowback gasconveniently comprises recycle gas, such as from conduit 41.

The amount of catalyst charged to the reactor initially is regulated,with reference to any preliminary treatment of the catalyst in thereactor and the gas velocity to be employed, whereby the upper level ofthe dense phase is substantially lower than the top of reactor 11.During the operation the accumulation of deposited reaction products onthe catalyst particles may cause an expansion of the dense phase and arise in the height of the dense phase. In certain of the operationsdiscussed hereinafter the dense phase became extended up into members 15and 16, and in other operations a portion of the catalyst was withdrawnto control the volume of the dense phase.

In the operation of the apparatus of the drawing, the desired quantityof powdered catalyst is introduced directly into the reactor through asuitable connection, not shown, in conduit 15. After any desiredpreliminary activation treatment, the temperature of the fluid in jacket25 is adjusted when necessary, by conventional heating or cooling meansor by controlling the pressure of jacket 25, to the temperature desiredto be maintained in jacket 2S during the reaction. After the catalystmass has reached the desired reaction temperature the introduction ofthe reaction mixture through pipe 13 is initiated. The reaction mixturemay be preheated by means not shown approximately to the reactiontemperature prior to its introduction through pipe 13 or the reactantsmay be heated to the reaction temperature through the passage thereofthrough that portion of pipe 13 which is' enclosed by jacket 25 and bycontact with the hot catalyst. It will be understood, furthermore, thatthe enclosure of pipe 13 in jacket 25 is not necessary to the inventionand that the reactants may be heated to the reaction ternperaure solelyby contact with hot catalyst. Generally, reactor 11 is maintained at asuperatmospheric pressure during both activation and hydrogenation.

Pipe 13 is dimensioned with respect to reactor 11 and the desiredsuperficial velocity whereby the linear velocity of the gases passingthrough pipe 13 is sufficiently high to prevent the passage of solidsdownwardly into pipe 13 against the incoming gas stream. A ball checkvalve, not shown, is provided to prevent solids from passing downwardlyout of the reactor when the gas stream is not being introduced into pipe13.

The reaction effluent from reactor 11 is removed therefrom througheither or both conduits 23 and 24 and passed by means of conduit 31 to aprimary condensation unit 32. Condensation unit 32 comprises a jacketedac cumulator in which steam is passed around the accumulator through ajacket to cool the reaction effluent to a temperature of about 300 F. atthe operating pressure existing in reactor 11. Cooling of the reactioneffluent at the operating pressure to about 300 F. condenses therelatively high molecular weight organic compounds and waxes and a smallamount of water, which products are removed from the condensation unit32 through conduit 33. Uncondensing vapors are removed from condensationunit 32 and passed through `a condenser 36 to accumulator 37.Condenser36` cools the reaction 37. The two liquid phases formed inaccumulator 37' comprise a heavy water-rich phase containing dissolvedoxygenated organic compounds and a lighter hydrocar-` bon-rich phasewhich also may contain some oxygenated organic compounds having morethan four carbon atoms per molecule. The two liquid phases are withdrawnfrom accumulator 37 through conduit 38 for subsequent recovery andpurilcation by conventional means not shown, such as by distillation andextraction. Uncondensed components of the reaction effluent comprisingunreacted hydrogen and carbon monoxide, methane and .carbon dioxide areremoved from accumulator 37 through conduit 39. These gases are recycledthrough conduit 41 to inlet conduit 13 of reactor 11 to supplement thefeed thereto and to alter the ratio of hydrogento carbon monoxide andthe carbon dioxide concentration in reactor 11. The presence of methane,excess hydrogen and diluents in the recycle stream serves to strip therelatively heavy organic compounds and waxes from the catalyst particlesin reactor 11 and is thus an aid in preventing settling of the fluid-bedof catalyst. v

In this apparatus operating runs were made to test the efficacy of theprocess in the treatment of a gas charge containing hydrogen and carbonmonoxide to convert these reactants to hydrocarbons and oxygenatedcompounds. In each operating run conditions were varied to test theeffect of various combinations of operative conditions. The results ofeach operating run are represented by the results observed during astabilized period of operation under a given combination of operatingconditions. The conditions of operation and the results obtained inthese operating runs are described below in the following examples.

1n the following more detailed description references to linear velocityin the reactor are based on the crosssectional area of the straightportion of the reactor, ignoring Ithe effect of the presence of thecatalyst. The inlet velocity is calculated from the gas rate enteringthe bottom of the reactor, with correctons for temperature and pressureexisting at the bottom of the reactor. The average superficial linearvelocityis calculated from the arithmetic average of the gas rate at thebottom of the reactor and at the 'top of the reactor. The lattervisarrived at by correcting the outlet gas volume for water andhydrocarbons condensed in the receivers, with cor# rections for pressureand average catalyst temperature. Contact times referred to below arethe superficial time5 in seconds, that the gas taken in passingthroughthe dense phase of the catalyst bed. It is calculated by dividing thedense bed height by the average superficial velocity.

EXAMPLE I iron. To prepare this material for use in this improvedprocessit was rst ground to a 6-20 mesh sizeand thensubjected to leaching withwater to remove lpotassium oxide. This treatment reduced the potassiumoxide content from 1.7 per cent to 0.55 per cent. The leached materialwas then dried at 210 F. land reduced in of hydrogen.

In the reduction treatment a heatedstream'of hydrogen was passed througha granular mass,I treated lto remove water formed bythe reductionreaction, and then reef circulated.` The temperature was raisedgradually and'l the reduction reaction was initiated at about 700-800f FThe temperature of the catalyst mass was then raised to about 1215"a F.in 2 hours while continuing the ow v of the hydrogen stream, During thenext 4 hours the temperature was raised to approximately 1285 F.du'ringwhich time the reduction was substantially completed, as evidenced bythe practical cessation of waterformation.

The reduced catalyst was ground in an atmosphere of COz, l-`1rst in ahand grinder and then in a ball mill, to

produce a powder having the following screen and roller analyses Rolleranalysis Particle size: Percent 0 10 Microns 11.0 10-20 16.4 20-40 20.640-60 32.2 60+ 19.8

Screen analysis U. S. Standard sieve: Percent `+40 mesh Trace 40-60Trace 60-80 0.5 -100 `0.5 1GO-120 Trace 120-140 TraceI 140-200 13.5ZOO-Pan 84.5

11,316 grams of catalyst thus prepared were charged;y into reactor lthrough an inlet (not shown) in section 5. f During this operation thecatalyst was maintained in the i atmosphere of CO2 and a Small stream of1 or 2'cu. ft. of CO2 per hour was passed upwardly throughreactor- 1 toprevent packing `of the catalyst. After the catalyst g was charged toreactor 1 the CO2 stream was replaced g with a stream of hydrogenwhichwas passed upwardly through reactor 1 at the rate of 15 to 20 cu. ft.per hour. The reactor was then heated externally while hydrogen l Vwaspassed upwardly through the reactor at this rate. When a temperature of530 F. was reached the hydrogen; stream was replacedby a stream ofsynthesis gas consist-f ing essentially of H2 and 4CO in the ratio` ofaboutl2:1..`

v,Thesynthesisgas-was passed upwardly through reactor` 1 at the rate of32-46cu. ft. per hour. At the Sametime:

the outlet pressure on the reactor was increased to 15; pounds. After 3hours at this condition the flow rate was raised to 60 cu. ft. per hourand the pressure was raised to 30 pounds. After 5 hours longer the llowrate was raised to cu. ft. per hour and the pressure was raised to 60pounds. After 5 hours of operation at the lastmentioned condition theflow rate was increased to cu. ft. per hour and the pressure was raised'to 100 pounds. At that condition the desired conversion of H2 and CO tohydrocarbons was soon achieved andpsubv sequently the pressure wasreduced to 80 pounds to con- 'trol the rate of reaction.

yBecause ofthe extreme rapidity of the strongly exo'- vthermic kreactionbetween Hz and CO relative to the rate a stream l ofl mixing at therather low linear velocities employed, the initial period of operationmay be considered as a catalyst activation conditioning, or inductionperiod. Duringthe first several days of operation temperature conditionswereobserved to be, somewhat derent from those observed subsequently.Following this preliminarycond 12 stabilized operation thereA were shortperiods of operation in which. operatingV conditions were being changed.Results observed duringy theseperiods of unsettled operation are notpresented, as they would. be Without significance.

o: In eiect, therefore, each of the periods of operation for which datavare presented in Table I represents an independent run whose results arecomparable with the results of the other run's, except for changes inthe condition of the catalyst. The superficial contact times employedin' 10 these periods' ranged from 35 to 95 seconds. The data TABLE I A BC D E F G H .T K L M N Operating Conditions:

Reactor Temp., Ave., F.- Y

12 ft. above pipe 3.. 535 528 524 531 540y 540 543 580 649 611i 600 10.5ft. above pipe 3 550 545 538 549 543. 531 532 545 535 572 649 599 5968.5 it. above pipe 3. 562 556 562 581 558 552 546 559 547 570 650 615609 6.5 it. aboveppe 3. 566 578 572 587 594 587 549 567 554 568 651 614610 4.5 it. above pipe 3- 587 598 628 614 620 613 581 582 566 572 664622 617 2.5 ft. above pipe 3. 584 592 606 588 580 562 565 585 569Y 569661 620 618 1.5 ft..above pipe 3- 561 572 612 579 546 538 548 588 572570 662 622 619 0.5 it. above pipe 3. 479 364 372 403 484 512v 519 587571 569 644 616 616 Feed Gas Temp., F-- 438 441 425 435 436v 444'v 449 v499y 507 568 I 616 544 505 Reactor Outlet Pr., p. S. 98 y 80 81 80 81 8180 81 149 52r 49 80 80 Gas Throughputs, s. c. t./hr. y

Gas Entering Catalyst Bed. 132. 2 146. 4 133. 7 162. 7 215. 1 250.1 293.3 311 406. 4 291. 7 204. 8 292. 7 194.8 Normally gaseous componentsleaving reactor. .7 82. 7 73.0 87. 8 107. 120. 9 168. 5 155. 5 186. 2214.0 96. 9 153. 5 94. 8 Blow Back t Filter. 29. 7 33.4 17. 9 25.0 24. 225. 1 31. 7 31.2 18. 4 22. 5 22.8 26. 6 17. 9 Analysis-Gas EnteringPercent- Hydrogen 62. 63. 9 63.8 63.1 62.6 62.4 62.9 61.6 62.9 62.1 62.654. 7 Carbon Monoxide. 33.0 30.4 31.2.: 33.2` 33.7 34.4 3155 31.2 29.632.6 31.0 38.8 Carbon Dioxide 1. 6 1. 6 2.1 1. 4 0: 7 1.1' 1. 8 2. 7 2.5 1. 8 1. 2 0. 9 Hydrocarbons and 111er 2. 9 4. 1 2. 9 2. 3 3. 0 2. 1 3.8 4. 5 5. 0 3. 5 5. 2 5. 6 HzZCO-Gas Entering Bed.. 1.89 2.10 2. 04 1.901.86 1.181 2. 00 1. 97 2.12 1. 91 2.02 1. 41 Inlet Velocity, 0. 49 0. 430. 54 0. 76V 0. 90 1. 06 1. 21 0. 96 1. 45 1. 05 1.18 0.81 Yi 1gO/HL/Lb.Fe, S. 0.1 2. 1 1.8 2.2 3. 7 4. 5 5.3 5.2 6. 7 4.6 3.6 4. 8 4.0

Vol. Percent Contraetlon 45.1 41.9 44.1 44. 5 48.9 50.7 41.2. 50.0 52.824.6 51.3 46.1 49.8 Cys, ce/cu. meter of feed gas i 32 35 23 34 44 Crs,ca /eu. meter of feed gas 8v 5 14 12 13 Crs, eeleu. meter ot feed gas 66 2 11 11 Light Naphtha, cc./cu. meter of feed gus 11 20 15 20 19 23 1319 26 27 Heavy Oil, cc./cl1. meter 0f feed gas. 54 48 41 42 42 50 42 4762 13 19 20 49 Tot. Liq. Hydrocarbons, cc./cu.

meter of feed gas 103 116 100 127 144 Water,'cc./cu. meter 0f feedgas--." 50 40 69 64 58 62 42 67 79 21 74 70 59 Tot. Liq. Hydrocarbons,Ga1s./

Day/Lb. Fe 0.21 o. 2s o. 2s o. 38 o. 27 Oxygenated Compds in Water,

cc./cu. meter oi feed gas 7 5 7 8 7 7 5 6 10 2 5 5 3 Percent H1Disappearance... 47. 7 48.2 52.0 51. 7 55. 5 56. 8 44.8 53. 8 62.0 30. 763. 5 57. 3 65.9 Percent CO Diseppearan.. 190 100 100 100- 100 100 85.6100 100 57.9 100 98.8 100 1:11:00 Reaction Ratio 0. 91 0.91 1.1091.05 1. 05 1.05 .95 1.10 1.22 1.12 1.21 1.17 0.93 CO Distribution- MolPercent to CO1 39.0 34.2 37. 9 37. 3 M01 Percent to C1214.-. 12. 2 9. 210. 1 10. 4 Mol Percent to 02's 6.8 8. 0v 6. 4 8.3 MolPercent to Css andheavier. 38. 6 44. 0 42. 6 42. 7 Mol Percent to Oxygenated Compds 3.44.6 3. 0 2.4 1. 3 Mol Percent to Crs and heavier- 26. 3 31. 7 29. 4Heavy Oil Inspections:

' Gravity, A. P. I 51.8 53.4 50.0

ASTM Distiilation, F

I. B. P 121 132 178 179 196 197 264 264 333v 331 445 445 644 670 738crkd. E. B. P crkd. at 94. 5 Mol Percent Mono-oletlns 64.8 AdsorherNaphtha Inspections:

Gravity, A. P. I 80.3 83.2 ASTM Distiliation, F.-

I. B. P 85 S0 5% 98 86 10%-. 102 89 115 97 131 112 156 132 219 198V 251'250 E. B. P 292 307 Reid Vapor Pressure, p. s. i 15.0 20. 2 Mol PercentMono-olens.- 70.2 78.8 Hours On Condition 24 17 TotalOpereting Hours.214 231 266. 5

The Ydata in Table I arearranged to present the results observed-fin 13periods `oli-'stabilized operation'durin'g this# operating .runflBeforeuandaftereach'foffthese periods of.'

inf1inef20arebased onf the total quantity of iron in the catalystinitially charged to the reactor.

W The eiective "D charge'f'rateswouliftlrerefore, `befsomewhat higherthan the figures given in line 20, since some of the catalyst would beretained on the sloping surfaces .of the apparatus yat 4 and 6 and outo' effective contact with the stream of reactants. Likewise some of thecatalyst forms a permanent mat on the filter surface. At the beginningof the operation of Table I the aeration of thecatalyst bed resulted inan averagedensity of *thev pseudofliquid lluidized dense phase of over 100 pounds percubicfoot and the upper level of the dense phase wasapproximately feet above pipe 3. However, accumulations of carbonaceousdeposit on the catalyst particles, which doubled the weight of thecatalyst mass and reduced the density of the catalyst particles, reducedthe density of the iluid bed whereby the upper level of the dense phaserose substantially higher than 10 feet above pipe 3 to levels inmanifold 6. The density of the dense phase was reduced also bythesubsequent increases in the inlet velocity of the gas stream. The'combination of these,

effects reduced the density of the dense phase Ato about 40 pounds perfour-fold. l Y.

During the whole operation of Table I the flow of the reaction mixtureout of the reactor was alternated bef tween filter 9 and lter 10 every15 minutes, and the oli-stream filter was blown back with feed gas atthe rate necessary to clear the lter of adhering the catalyst.

During the run of Table I the reaction products were recovered for themost part by cooling the reaction mixl ture to room temperature, orlower, to obtain a condensate, and then passing the remaining gasthrough an adsorbent. The condensate comprised bothlheavy oil `and waterproduct fractions. The heavy oil fraction contained a small quantity ofoxygenated compounds and the water product fraction containedsubstantial amounts of Oxy# genated compounds. The adsorbed product wasrecovered by steam distillation, which produced a light naphtha fractioncondensate water and a gas fraction. The condensate water containedadditional oxygenated compounds.l

cubic foot and increased the volume about The gas fraction was almostentirely hydrocarbons having 3, 4 or 5 carbon atoms per molecule. Theyields of the various fractions were determined by measurement of thecondensed product and by absorption and combustion analyses of the gasfrom the condenser. l,

After period L, hydrogen at a rate of 15 cu. ft. per

hour was substituted for synthesis gas. Simultaneously.

the pressure was reduced to one atmosphere and the temperature to 500 F.The hydrogen was kept in for nine hours, after which synthesis Wasresumed.

Immediately following period M and after 737'hours of operation, thereoccurred a brief period of operation at relatively high temperature.Following this high trern-` .f

perature operation, the catalyst was subjected to a reviviticationtreatment with hydrogen. In this treatment, the unit pressure wasreduced to 30 pounds per square inch and one hour later the synthesisgas was replaced with hydrogen at about 25 cubic feet per hour and thetemperature was reduced to about 500 F. After 4 hours, the temperaturewas raised to about' 615 F. over aperiod of 3 hours. The pressure wasmaintained at 30 pounds per square inch for 6 hours of the hydrogentreatment and then raised to 80 pounds per square inch for thelremaining 3 hours. v

Operating period N, which immediately followed the revivificationtreatment, may be referred to for an example of the results obtainedduring this operating run. In ythis period of 36 hours the Dowtherm A injacket 15 was maintained under a pressure of 20 pounds per square inchto produce a temperature in jacket 15 of 580 F. During this operationthe height of the dense bed was approximately 15.4 feet above pipe 3,whereby the upper level of the dense bed was located in manifold 6. Thecatalyst density in the dense phase was approximately 44 pounds percubic foot whereby the space velocity was approximately 160 volumes offeed gas per hour per volume of dense phase. The quality of the liquidproducts obtained is indicated in Table I. The gas tra@ tion obtainedfrom the adsorber comprised about 75 per cent olefins. The gas from thecondenser contained no carbon monoxide not attributable lto theblow-back gas. This indicated complete conversion of carbon monoxide inthe reactor to hydrocarbons, oxygenated compounds and carbon dioxide.The amounts of these products detected in the reaction product mixtureaccounted for 99 per cent of. the' carbon .monoxide charged to thereactor.

For an example ofthe quality of the product made duringvthe operatingrun of Table I, reference can be had -to rdeterminations made'o'n theproduct obtained in period H. These' determinations were made on aspecimen prepared by blendingy the light naphtha and condensed oil,debutanizing the blend, and distilling it to 300 F. end point. At thesame time a diesel oil boiling between338 F. and 650- F. was obtained.The raw gasoline thus obtained was 65.1 per cent of the total of -thelight naphtha and condensed oil. The diesel oil fraction accounted Afor31.5 l. per cent, and the remaining 3.4 percent `was material boilingabove 650 F. The raw gasoline fraction had an aniline point of 88 F., agravity of `73.1 API and a Reid vapor pressure of 8.5 pounds persquareinch. The diesel oil fraction had an ASTM pour point of `15 F. anda diesel index of 61.5. The

octane number determinations on the raw gasoline are given in theiirstcolumn of the followingtable.

Raw 0 lotal Gasoline ltfy Product Octane No. ASTM) 69 71 74 Octane No.ASTM)+3 cc. T 79 80 82 .Octane No. (CFR-R) 77 8O 84 Octane No. (CFR-R)+3cc. TEL 91 92 94 Ir the foregoing table the octane numbers given underthe heading 100% Pentane Recovery are based on blending, Awith therawgasoline, all the pentane-pentane vapor pressure of 10 pounds per squareinch.

A concentrate of low molecular weight oxygenated chemicals was obtainedby careful distillation of a blend of several aqueous fractions producedat 80 pounds per squareincl.

Originally, the water layer contained approximately 8 per centoxygenated compounds. The distillation was conducted in a batch columnequivalent to about 10 theoretical plates. Before commencing thedistillation,

ha small amount of caustic was added to the still pot to neutralize theorganic acids present.

Formaldehyde and acetaldehyde were present in the original Watersolution,

`but the amounts were very small and not determined.

The concentrate of oxygenated compounds from the primary distillationwas subsequently refractionated for identification purposes. Fiveprincipal cuts were obtained which were predominantly acetone, methylethyl ketone, ethanol, n-propanol, and n-butanol, results for whichappear below.

. Water Yield on y Vapor, Content Concenr-BlendlNumber Temp., PrincipalConstituent of Cut, trate, F. Wt. Per- Vol. Per

cent cent 131-132 Acetone O. 5 18 162-163 Methyl ethyl ketone.- 7. 6 7171-175 Ethal101 7. 6 30 187-190 n-Propanol 7. 7 15 1-`198-200 n-Butanol5 15 As the distillation progressed to'y af vapor temperature above 192F., a condensate consisting of two phases appeared. The upper and lowerphases were principally n-butanol and water, respectively.

EXAMPLE 11 The catalyst for use in this operation was prepared from thesame source material as the catalyst in Example I, and by the samegeneral-procedure. In this case, however, the catalyst was leachedsufliciently to reduce the potassium oxide content from 1.7y to 0.41 percent. The leached granular material was dried at 200 F. overnight andthen reduced in a stream ofhydrogen in the general manner described inExample I. Reduction of the catalyst was initiated at about 700 F.thereafter and the temperature of the catalyst mass was raised to about1350 F. in 4 hours, while continuing the ow of the hydrogen stream. Thiscondition was maintained for 2 hours longer, during whichtirr'ie thereductionl was" substantially completed. The reduced mass was thencooledto room temperature in the hydrogen atmosphere.

The reduced catalyst was then ground, first in a hand mill and then in aball mill, to the de'sired'degree of lineness. Throughout this periodthe catalyst` was not permitted to come in contact with air, theVgrinding operations being conducted in an atmosphere of'COz. Thecatalyst powder had the following screen and roller analyses:

Roller analysis particle size; pei-cent 35 This operation was continuedfor a total run length of 10 microns 17 1166 hours, after which theoperation was terminated -20 19,5 voluntarily to free the apparatus foranother operation. -40 24 During this period the superficial gas Contacttimes were 40-601 32. varied between 14 and 35 seconds. The resultsobserved 60-l 7.5 40 inpenodsof stabilized operation are set forth inTable Il.

TABLE II A B C D E F G H J K L M N P Q. R S

Operating Conditions:

Reactor Temp., Ave., F.- v

12 fr. above pipe 8. 540 549 548 562 588 585 577 588 523 500 506 539 557575 555 580 580 10.5 n. above pipe 8- 541 548 548 551 575 580 574 548525 v 510 528 547 552 577 558 548 548 8.5 fbabove pipe 8. 545 558 555567 575 581 576 550 542 586 556 574 587 598 584 575 570 6.5 it. ebovepipe 8.- 550 559 561 575 578 585 588 569 554 546 558 598 508 508 508 591587 4.5 it. above pipe 8-. 560 575 582 594 587 594 590 579 551 558 587511 518 618 518 510 505 2515.561575 pipe 8-- 571 594 508 507 588 595 589579 558 555 504 612y 514 618 508 521 517 1.511.5bove pipo 575 598 514509 585 594 587 577 554 555 605 606 612 620 507 528 519 0.5 it. abovepipe 8 578 582 592 598 580 591 576 578 544 589 587 588 598 610 578 517515 Feed Gas Temp., F 450 870 408 512 50i 508 525 508 507 545 459 381458 575 481 552 545 Reactor outlet pressure,p.s.1.g- 102 148 150 245 250245 151 150 248 250 152 251 249 249 249 250 255 Gas Throughputs, s. c.L/hr.-

Fresh Feed 220.5 275.2 402.4 .404.0 829.4 851.5 281.5 277.8 804.8 220.8880.7 425.5 822.8 467.8 285.7 257.0 210.7 Ges Entering Catalyst Bed.-810.2 401.1 402.4 629.9 570.3 555.7 404.9 408.9 586.7 585.0 880.7 587.8856.6 901.8 585.2 424.0 410.7 Normally gaseous componentsieeving reactor155.1 209.7 175.8 848.2 448.4 407.5 280.52155 400.0 475.0 202.7 579.4550.5 527.5 418.8 275.1 259.9 Net Product Ges 98.8 87.0 175.8 141.7109.5 104.4 100.8 90.1 85.5 50.4 202.7 188.2 105.2 198.5 109.8 58.9 70.0Blow Beek to Filter-- 9.0 18.5 28. 6 81.5 81.1 28.5 28.7 16.0 10.4 10.915. 8 14. 9 15.2 17.0 12. 8 19.2 5. 5 Recycled Gas: Fresh Feed .3 .5 0.5 1.0 .8 .4 .4 .9 1.8 0 1.0 1.6 .9 1.0 1.0 .9 Analysis-Gas EnteringBed,

Mol Percent- Y Hydrogen 50.5 56.5 59.7 58.1 55.2 54.1 59.0 59.5 55.250.4 `58.8 52.8 88.5 50.8 45.8 49.1 54.1 Carbon Monoxide-- 24.1 24.185.0r 22.7 18.7 21.0 25.1 23.5 19.8 14.9 25.8 19.0 19.5 24.2 25.9 18.816.5 Carbon D1oride. 7.4 9.8 1.1 9.0 18.1 11.9 7.5 5.8 11.2 18.7 2.912.5 22.1 11.9 14.8 12.4 8.8 Hydrocarbons 7.8 10.1 4.2 10.2 18.0 18.08.8 10.1 12.8 21.0 8.5 15.2 19.8 18.5 11.5 20.2 21.1 Heco- Ges EnteringBed 2. 5 2.8 1. 7 v 2. 5 2. 9 2.5 2.8 2. 5 2. 8 8. 4 2. 7 2. 8 2.02.1 1. 7 2.7 8.8 Fresh Feed- Total Hi and CO M01 Per- Y.

een 95.1 95.8 94.7 948 948v 95.4 94.7 95.0 95.8 94.9 98.5y 95.9 98.994.1 95.4 98.9 91.9 Hi 1.9 1.8 1.7 1.9 1.8 1.7 1.8v 1.9. 1.9 2.0 2.7 1.81.8 1.5 1.8 1.9 2.1 FreshFeed/Hr.v e,s.o. 11.9 14.9 21.5 21.7 18.5 19.915.8 16.8 18.4 18.5 l88.1 y89.0 82.5 51.8 88.5 27.2 27.7Tote1Ges/Hr./Lb. Fes. o. 15.7 21.5 21.5 88.9 87.9 87.1 28.4 28.8 88.889.2 88.1 79.5 87.4 99.0 68.8 55.8 540 CO in Total Gas Charge/HL] Y 57.758.5 56.2 55.0 56.8 70.2I 54.2 57.51715 :72.5 45.7 68.0 57.0 58.5 51.571.6 55.8 80 25 38 85 85 88 86 84 82 88 80 48 85 84 42 88 85 4 11 8 1618 19 18 18 16 18 12 15 18 22 20 12 i0 89 16 14 18 12 21 15 20 18 ii 8 520 10 14 22 15 29 85y 28 85 8 87 88 29 28 88 28 i5 84 25 28 i8 25. 42 5154 58 68 45 40 52 55 85 27 58 89 88 88 54 28 15 Screen analysis U.Standard sieve: Percent 1r`1esh Trace 40`60 Trace l80 Trace -100 Trace-120 Trace -140 Trace -200 5 200-Pan 93.5

94,080 grams of this catalyst were then charged into reactor 1 lby theprocedure described in Example I. After startingthe passage of hydrogenthrough the reactor at the rate of 15p-20 cubic feet per hour the outletpressure on the reactorwas then raised to 80 pounds and the Ltemperaturein the reactor was raised to approximately 450 F. by means of theheating coils around jacket` 15. The hydrogen flow rate wasthenincreased to 5,0 cubic feet per hour and the temperature was thenraised to 500 F. Then the hydrogen stream was replaced wth a stream ofsynthesis gas consisting essentially of Hzfand CO in the ratio of 2:1.The synthesis gas wasV passedupwardly through the reactor at the rate of140 cubic feet per hour. After one hour the temperature was raised to550 F. and the ow rate was increased to 200 cubic feet per hour. lAfter5 hours longer the ternper'ature was raised to 620 F. and after 3 hoursoperatioii'at 620 F. the ilowrate was increased to 325 cubic feet perhour. At that'point conversion of the H2 and COftohydrocarbons started.and the temperature was reduced to' 600 F. Operation at these conditionswas continued for 35 hours longer, at which time the pressure Was raisedto 100 pounds.

A B O D E F G H J' K L M N P Q R S Yields (Based on Fresh Feed)-Oon.

Oxygenated Compounds, 00./011.

meter 14 14 13 17 18 15 12 14 18 20 1i 15 1i 13 13 14. 14 Total Liquid,oo./ou. meter 151 158 155 174 155 175 149 152 153 155 111 152 152 137155 153 127 Water Produced, oe./ou.meter-- 98 125 75 119 138 119 99 110125 125 75 115 80 55 58 -114 118 PercentHzDisappearance (overal1 orfresh feed basis) 54.4- 785 52.2 72.5 73.4 81.1 71.1 74.8 78.8 85.5 48.478.0 84.2 71.5 77.1 84.3 78.3 Percent CO Dlsappearance 1 (over-511 orfresh feed bes1s) 1 00 100 100 100 100 100 100 100 99.5 100 95.0 10095.4A 92.5 93.2 100 .100 112:00 Reaction Ratio 1.24 1.42 1.05 1.35 1.311.4i 1.30 1.43 1.50 i. 57 1.38 1.37 1.13 1.13 1.12 1.59 1.58 CODistribution- Moi Percent to Coe 27.1 24.9 32.1 19.3 20.3 21.3 22.9 23.921.7 19.8 23.0 27.3 34.2 33.0 31.8 21.5 18.9 MolPeroent to OH. 10.8 11.18.7 9.5 9.8 10.3 10.2 5.7 4.9 9.5 15.2 10.5 9.0 10.5 11.2 11.8 15.5MoiPeroent to Ci's 5.5 4.7' 5.5 8.5 9.2 8.3 9.1 8.7 8.4 10.5 `11.3 l9.18.4 8.8 9.0 9.3 10.5 Mol Percent to Ces and heavier 51.5 55.5 49.1 58.155.4 55.9 54.4 57.7 59.7 53.7 45.8 50.3 45.5 43.9 44.9 53.1 50.7 M01Percent to Oxygenated f Oompds 3.9 3.7 3.5 4.5 5.3 4.2. 3.4 4.0 5.5 5.83.7 2.8 2.8. 3.7 3.1 v4.3 4.3 Mol Percent; to 's and heavier 42.0 44837.7 42.5 37.5 39.5 38.3 42.2 43.7 34.9 29.8 32.1 32.8 25.8 28.4 38.533.5 MolPeroent to H40 44.8 49.1 35.5 59.5 57.5 55.8 53.5 51.4 54.8 58.053.5 44.9 32.0 33.4 35.8 55.8 50.8 Heavy O11 Inspections:

Gravity, .4.P.I 53.5 55.5 57.2 58.0 57.9 57.9 55.3 55.9 y55.4 58.1 59.953.5 57.2 54.5 54.8 53.1 59.2 ASTM Distillatlon, F I. B.P 131 114 105110 105 105 115 115 117 118 113 94 107 112 114 97 108 5% 187 l158 152145 144 144 174 154 153 119 140 175 154 128 152 209 182 154 157 154 154194 195 191 183 185 159 194 184 141 172 272 248 225 235 234 234 259 253255 253 229 188 239 258 -255 191 231 335 305 284 293 299 299 318 315 332335 283 250 305 335 330 248 297 430 392 380 390 393 393 418 408 534 441359 '330 400 440 434 334 355 510 548 554 585 581 581 519 580 548 707 513505 508 555 n 712 529 539 706 630 655 crkd. 702 702 700 688 crkd crkd.633 618 726 orkd crkd. 667 659 E. B.P 718 598 558 crkd. crkd. orkd. 702.540 528 orkd. crkd.' 575 MolPeroent Mono-olefin 75.8 75.0 75.9 75.0 73.573.5 71.9 54.5 55.9 54.2 59.8 55.1 58.9 59.0 57.5 53.4 50.1 Adsorber Naphtha Inspections:

Gravity, A.1 .I 75.9 79.5 82.0 78.7 70.4 75.5 79.4 (81.0) 79.8 78.9 80.280.8 75.3 75.4 75.9 74.7 78.8 ASTM Distinction, F.-

1.13.18 89 87 82 85 114 85 87 82 89 92 89 94 .90 94 92 93 90 5% 109 10092 98 144 104 97 95 101 104 101 105 104 112 107 105 102 10% 115 105 94108 155 110 101 98 107 11i 105 108 110 118 113 113 107 30%-- 135 118 103124 179 130 Y113 110 121 125 117 117 129 135 128 138 121 50% 150 133 115142 202 152 134 122 135 145 132 127l 153 155. 145 173 138 70%-- 181 155135 159 227 184 155 145 159 171 150 145 f 171 -183 175 223 155 90%" 219202 188 218 255 242 228 203 207 215 221 197 253 238 240 275 230 95% 235228 225 241 255 254 238 244 243 241 255 243 291 275 284 3ii 275 13.8.13258 255 252 280 315 292 290 274 274 278 284 280 325 334 338 357 334 ReidVapor Pressure,p.s.i 12.5 15.5 17.5 13.2 5.5 12.9 15.0 15.8 13.7 18.013.8 13.8 12.4 10.1 13.1 v11.8 14.4 Moi Percent Mono-o1efins 77.0 78.178.8 75.4 54.0 75.5 77.2 73.8 72.3 70.4 55.9 (74) 72.8 72.2 53.4 58.354.5 Hours on Condition 43 48 44 45 24 48 58 48 72 50 52 34 48 24 24 1851 Toteioperating Hours 152 195 242 255 314 382 447 519 579 579 725 828940 988 1,052 1,114

The data in Table II are arranged to present the results 40 in the densephase decreased from. about 16' pounds in observed in 17 periods ofstabilized operation during the period A to about 5 pounds in period S.-The catalyst operating run. Throughout this operating run the freshdensity decreased from an initialvngure of over 80 pounds feed to theoperation contained hydrogen and CO in ratios per cubic foot to about40-pounds in period P. The varying from 1.3:1 tov2.7:1. In most of theperiods of density also was aiected-by changes in the velocity ofstabilized operation presented inw Table II suicient uri- 45 the gasespassing` through the' reactor. l converted gas was recycledto increasethe HzzCO ratio During the operation of Table II the ilow of the reacinthe gas entering the catalyst' bed substantially above tion mixture outof the reactor lwas alternated between the corresponding -gure in-thefresh feed. In the recy4 lter 9 and lter 10 every 15 minutes,andthe'ol-stream cling voperations .the vH2:CO ratio in gas entering thefilter was blown back with feed gasat-the-rates indicated catalyst bedvaried from 1.7:1 to 3.4:1. I n most cases 50 in Table II. f l i nounconverted C0 was observed in the product mixture Thereaction productswere recovered during the opwhen operating even at relatively high spacevelocities. eration of Table II by cooling the reaction mixture to Forexample, at a'charge rate as high as 15.1 cubic feet room temperature,or lower, to obtain a condensate and of carbon monoxide per hour perpound of'iron complete n thenpassing' vthe remaininggas throughar i vadsorbent-- disappearance of the CO wasobserved. At the'm'axim'um 55'Heavy- Aoil a-nd water product fractions were obtained charging rateemployed, 24.0 cubic feety of CO per hour from the condensate. The heavyoil. fraction contained,- Pel Pound 0f iron, only ,2811er Cent 0f the C0Charged t0 in addition, oxygenated compounds, such as butyl, amyl, thereactor was observed in the product mixture. This hexyl and heptyIalcohols. The water product fraction Ooonffed n Period P, during Which.time the tofl gas contained substantial amounts of oxygenatedcompounds, charge entering the bed was introduced at the rate of 99.0 60such as ethyl, propyl and butyl alcohol, acetone and Cubic feet Per hourPer Pollnd of fon This corresponded methyl ethyl ketone; The adsorbedproduct was recovo a space Velocity of 11,915 Volumes of gas enteringthe ered by steam distillation, which produced a light naphtha catalystbed'per hour per cubic foot of dense catalyst phase. fraction condensatewater and a- 'gasvfractiorr Thecon- During the operating fun *of TableII the Clnanly' of densate water yielded additional oxygenatedcompoundsiron CatalYS-n the reactor Waslednoed Peflodlcauy 11.1" 65 Thegas fraction was almost entirely hydrocarbons having the operation inorder to obtain samples for analysis, ing 3, 4 or 5 carbon atomsperlmolecul The amounts flmd m 9rd to reduce the -V Chim? 'of the-cafalyS-t mass of Vother compounds in the reaction product mixturelwas. m he reactor' -Th accumulatin of 'carbmaceous de' determinedbyabsorption, combustion and mass specposits on the catalyst'particlesincreased the volume of trometry l u ro ded so v .'1 tcllsgptllelieteagseogrs tegver 14 70 Forty-one hours after synthesis gaswasintroducedpto ri dof feet above pipe 3.'.to a level in manifold 6. Inorder to @e Systifll hyftll'ogen tllva addsed fiori a rSc/rtthpeeSeral.- maintain all of the dense phase in reactor 11 proper, and timet? 1 We e syn esl] ga Q mp f. h .d g th in order to permit operation at.relatively high space oPenanon- Inst befofe e a lflon o Y regen evelocities, the quantity of iron catalyst in thereactor was maXnnumtemperature 1n the CetelysV bed Wes on the 7'5 order of 690 F. Afterperiod G of the operation ofY progressively reduced whereby the amountof iron catalyst Table II the feed to the operation was changed tosubstantially pure hydrogen. At the same time the Dowtherm was removedlfrom jacket 13 and the reaction chamber was heated by means of theexternal heating coils to a higher temperature to eiect reduction of thecatalyst by means of the hydrogen. In this operation the hydrogen waspassed through the reactor at flow rates varying from 18 to 37 cubicfeet per hour and the temperature of the catalyst was raised in about I6hours to an average temperature of approximately 930 F. This conditionwas maintained, with a maximum catalyst temperature of 972 F., forapproximately 10 hours, after which time the temperature of the catalystwas reduced in 7 more hours to approximately 430 F. The temperature wassubsequently raised to about SOO-550 F. and maintained for a short time.

The pressure during the hydrogen treatment was held at atmospheric for.27 hours with hydrogen once through, then it was raised to 150 poundsper square inch and then hydrogen recirculation was started. Thefollowing hydrogen rates entering the reactor were used:

The total length of hydrogen treatment was 40 hours. Thereafter, thefeed to the reactor was changed to the synthesis gas mixture and the`conversion reaction Was resumed. Prior `to this reduction treatment ofthe catalyst the contact mass contained 0.195 pound of carbon, 0.268pound of oxygen, and 0.091 pound of oil and wax, per pound of iron.After the treatment the contact mass contained 0.170 pound of carbon,0.070 pound of oxygen, and 0.001 pound of oil and Wax., Per pound ofiron.

After this reduction treatment of the catalyst and after a short periodof variable conditions period H of the operating run was started. Theimprovement in activity of the catalyst following the reductiontreatment isshown 'by `a comparison of the data .of periods G and H,`which 'show that a Alower temperature after reduction provided the sameIra'te of 4conversion of `carbon monoxide as was reached previously atYa higher tempera* ture. 'These data show also 4an :improved vyieldofoil and reduced production of hydrocarbon gases.

Most of `the operating periods of the operating run of 'Table IIinvolved recycling of unconverted gas at 1ra tios of recycled gas tofresh feed varying from .3:10 to 1.181.144. .The recycling `operationsordinarily linvolved relatively high HzzCO ratios in 'the gas 'enteringthe catalyst bed. However, the `data indicate lthat, other conditionsbeing equal, and at the same HzzCO ratio, the operating runs underrecycling conditions produced substantially more foil per unit quantityof fresh feed.

One of the -benecial effects of recycling lies in the great improvementin the selectivity of the synthesis reaction. This point is clearlydemonstrated by `comparison of periods selected lafter 828 and 1114hours of operation.- Recycling was `employed in each case; lhowever, inthe former period the H2:CO ratio entering the reactors was a-nd in thelatter case, 3.3. A few salient features of each of `these tests arelisted below:

Zt) Thus the conversion to carbon dioxide, methane, ethlene, and ethanewas materially lower at Condition S than at Condition N.

In the 250 p. s. i. g. runs D, E, F, I, K, R and S of Table II, a majorproportion of the oxygen of the carbon monoxide reactant was convertedto Water and a minor proportion to carbon dioxide. In these 250 p. s. i.g. runs the minimum space velocity was about 6 standard cubic feet ofcarbon monoxide and about 34 standard cubic feet of total feed gas, perhour, per pound of catalyst. The minimum carbon dioxide content of thetotal feed was 8 volume per cent and the minimum H2:CO ratio in theinlet gas was 2.5.

In runs C and L of Table II, no recycle was effected. In runs A, B, C,G, H, and L, the pressure was about pounds per square inch gage orlower. It should be noted that -in all of the runs of Table I, norecycle and low pressures were employed. In runs B, C, N, P and Q ofTable II, the low water production was attributed, at least in part, tothe low Hz'rCO ratio in the inlet gas. Generally, the use of lowpressures resulted in low throughputs. The runs without recycle resultedin most instances in low HzzCO ratios and low concentrations of `CO2 -inthe reaction zone.

In the runs of Table II in which a major proportion of the oxygen of thecarbon monoxide was converted to water, the reaction mol ratio ofhydrogen to carbon monoxide Was at least 1.3.

Varying amounts of hydrocarbons and oxygenated chemicals were recycledwith the fresh feed to the synthesis reactor. The recycle stream wastaken after the product gas had passed through the secondary receiverwhich was ordinarily maintained, during the recycling tests, vatapproximately l40 F. and reaction pressure. During some of these.operations the recycle gas was passed through the Vcharcoal adsorberand stripped of the oxygenated compounds present and of all 'but lthe4very lightest hydrocarbons. In other operations the stream was notadsorbed and contained 'light oxygenated materials and hydrocarbons.,predominantly olens, through Cin. The .recycle stream was taken justafter the secondary receiver and after passing ythrough or -by-passingthe adsorbers .the .pressure was 4released and the gas fed tothe suctionof lthe lsynthesis compressor. When the recycle was passing through theadsorber -the product gas 'was merely vented to the atmosphere aftermetering and `sampling whereas during a number of the operations wherethe recycle was not adsorbed the adsorber was .on the product stream.When conditions A, B, D, F, G, .'H, I, Pand Q were obtained the recyclegas was taken -after passage over activated carbon, and when conditions'K and Svwere obtained the recycle was passed over activated carbon forthe greater part of the tests.

At the beginning of the operating run of Table II the reactor was.filled with catalyst to .a catalyst bed yheight of about 1-0 feet abovel-pipe 3. At an average superficial velocity yof about 0.76 feet ,per:second the dense v.phase had, Ain the lower A.portion thereof, a.density yof about .83 pounds per vcubic foot. As the yoperationproceeded the iron vbecame partial-ly oxidized .and the catalyst also.accumulated carbon .and deposits of oil and wax. After 586 hoursoperation a sample of the catalyst was withdrawn and aerated -with.inert gas at .1.2 .feet per second. The catalyst density was 45:0:pounds per cubic rfoot. A simil-ar test at l1'6`6 hours indicated adensity of 451:5 pounds :per cubic foot. 'Changes in lthe composition ofthe catalyst during the run are indicated below in 'the -tableheadedCatalyst Composition.

Catalyst composition Total Hours On Stream 245 383 383 445 588 940 1,114 1, 166 Hours After Ha Treat 0 62 203 560 7 780 Catalyst Analysis:Carbon, Wt. Percent.. 0 12.4 12.4 13. 4 15.1 1 .6 18.3 20. 9 23.2Oil-l-Wax 0 4. 1 5. 8 0. 1 1.8 .0 18. 9 18. 6 15. 6 Fe 93 63. 8 68. 779. 1 67. 1 5 8 47. 4 45. 0 46. 5 A1203 4 2O 0. 5 Iron Distribution:

Oil, Wax and C Free- Fe 89. 1 37. 2 46.0 Fe Oxides. 5. 0 57. 1 37. 0Lbs. C. 100 Lbs. Fe-- 0 19. 4 19. 5 16. 9 22. 5 22. 7 38. 6 46. 5 50.0Lbs. Cai7./100 Lbs. Fehn.' 107 157 157 126 149 167 211 222 215 Theeffect of the intermediate hydrogenation treatment of the catalyst isindicated in the two columns at 383 hours on stream. The specific effectof this treatment on the various ingredients of the catalyst has beenindicated above. The over-all effect can be seen in the foregoing tablein the reduction of the weight of catalyst per 100 pounds of iron from157 to 126.

Oxygenated chemicals were recovered from synthol oil produced at 250pounds per square inch by means of thefollowing procedure. The oilproduct was -i'rst caustic-washed to remove organic acids andsubsequently treated with a solvent to obtain the remainingoxy-chemicals. The condensed oil fraction subjected to the extractionstep was found to contain 6.25 weight per cent of oxygenated chemicals.Since a small amount of carbonyl compounds-aldehydes and ketones-wasproduced, this extract material was hydrogenated and then redistilled.The distillation indicated a distribution of alcohols in thehydrogenated extract from the oil shown below.

Percent Ethyl alcohol 8 Propyl alcohol 9 Butyl alcohol 13 Amyl alcohol22 Hexyl alcohol 22 Heptyl alcohol 13 Higher alcohols 13 The alcoholsproduced in this operation were principally straight chain.

A breakdown of the principal acid-free Oxy-chemicals from thisoperation, including those chemicals recovered in the water product, ispresented in the following tabulation:

Volume percent A small amount of carbonyl compounds recovered with thesynthol oil has not been included.

EXAMPLE III This example was a stabilized period of operation in arelatively long operating run carried out in the reactor whoseconstruction and operation are similar to that shown in the drawing anddescribed in Example I. The catalyst, employed in a finely dividedcondition, consisted principally of iron and contained 0.70 weight percent KzO. The average operating temperature in the reactor during thisoperating period was 551 F. 'Ihe fresh feed to the operation was 90 percent H2 and CO, the remainder being CO2, N2 and hydrocarbon gases. Inthe fresh feed the H2:CO ratio was 2:1. The fresh feed was charged to`the operation at the rate of 3.8 standard cubic feet per hour per poundof the original reduced catalyst present in the reactor. The tail gasfrom the operation was recycled to the reaction zone in a ratio ofrecycled gas to fresh feed of 5.2:l. As the result of recycle the ave'-rage space velocity was equivalent to at least 5 standard cubic feet ofCO, per hour, per pound of catalyst. During the operation nearly all theCO was converted to products other than CO2, a major proportion beingconverted to water. Consequently, the HzzCO ratio in the total chargegas was 5:1 to 10: 1, and the charge gas contained at least 6 mol percent carbon dioxide.

Under these conditions of operation the total liquid hydrocarbonproduction (including oil-soluble organic compounds) equaled 104 cc. percubic meter of fresh feed gas. This product was obtained by condensationat ice-water temperature and 250 pounds per square inch. This producthad an A. P. I. gravity of 56.1 and contained 69.8 mol per centmono-olens. At the same time the production of water (includingwater-soluble organic compounds) equaled 198 cc. per cubic meter offresh feed gas. The size of the water production indicated clearly thatthe oxygen eliminated from the system, in forms other than organiccompounds, was being eliminated largely as H2O, rather than CO2.

EXAMPLE IV This was a period of a relatively long operating run carriedout in a reactor generally similar to the reactor employed in Example Iand under generally similar conditions. The finely divided iron catalystcontained, per part of iron, 0.01 part A1203, 0.011 part TiOz, 0.008part SiO-2, and 0.014 part K2O. In this operating period the temperaturevaried from 610 F. near the inlet to 597 F. near the outlet, and apressure of 250 pounds was maintained on the reactor. Fresh feed wascharged to the operation at the rate of 45.3 cubic feet per hour perpound of iron in the catalyst in the reactor. Tail gas was recycled tothe operation in a ratio of recycled gas to fresh feed of 2.8:1. Thefresh feed gas contained 75.2 mol per cent H2 and 18.8 per cent CO, theremainder being CO2 and hydrocarbon. The total gas charge, i11- cludingfresh feed and recycled gas, contained 75.1 mol per cent H2 and 11.4 molper cent CO, the remainder being substantially all CO2 and less than 4mol per cent light hydrocarbon gases. Under these conditions 74.7 percent of the CO was reacted, including 2.3 per cent converted to CO2.This operation produced total liquid hydrocarbons to the extent of 63cc. per cubic meter of fresh feed gas and 21 cc. of oxygenated compoundsper cubic meter of fresh feed gas.

The selectivity of the operation was excellent when these conditionswere employed, e. g., 66.5 per cent of the carbon monoxide when reactedwas converted to oil (propylene and higher hydrocarbons) and oxygenatedchemicals, and only 3.3 per cent of the reacted carbon monoxide wasconverted to carbon dioxide. In spite of the greater hydrogenatingactivity of the catalyst as a result of the relatively highconcentration of hydrogen in the reaction zone the C2 fraction 86 percent propylene, and the Ct fraction 84 per cent butylenes. The oil whichwas condensed from the reactor eiiiuent at ice-water temperature andoperating pressure contained nearly 30 per cent oxygenated chemicals,largely alcohols and acids.

Having described our invention, we claim:

ln a process for the hydrogenation of carbon monoxide in which a gaseousreaction mixture comprising hydrogen and carbon monoxide is passedthrough a reaction zone containing a finely divided iron catalyst at avelocity effective to suspend said catalyst in said gaseous mixture insaid reaction zone under conditions such that hydrogen and carbonmonoxide are reacted to produce organic compounds as products of theprocess, the improvement during the synthesis proper which comprisesintroducing into said reaction zone a fresh feed gas containing hydrogenand carbon monoxide in a mol ratio in excess of the mol ratio in whichthese components are reacted and containing a relatively small amount ofcarbon dioxide, passing the gaseous reaction mixture through saidreaction zone at a rate equivalent to at least 30 standard cubic feet oftotal gas and between about 4 and about 15 standard cubic feet of carbonmonoxide, per hour per pound of iron catalyst in the reaction zone,maintaining a temperature of reaction between about 550 F. and about 750F., a reaction pressure between about 150 and about 600 pounds persquare inch gage and a contact time between the reaction mixture andcatalyst in the reaction zone of between about 8 and about 35 seconds,converting over-all a major proportion of the hydrogen and carbonmonoxide in a mol ratio of at least 1.3 to products of reactionlcomprising organic compounds having at least `one carbon atom permolecule and water, whereby under the conditions of operation a majorproportion of the oxygen of the carbon monoxide is converted to Waterand not more than a minor proportion is converted to carbon dioxide,withdrawing a gaseous eiuent from said reaction zone, cooling saidgaseous effluent to condense water, separating condensate cornprisingwater from an uncondensed portion of said eluent comprising hydrogen andcarbon dioxide, and recycling said uncondensed portion of said effluentcomprising hydrogen andv carbon dioxide to the inlet of said reactionzone in a'volumetric ratio of recycle gas to 'fresh feed gas of betweenabout '0.5 :1 and about 5:1 such that the total feed gas to saidreaction zone contains hydrogen and carbon monoxide in a mol ratiogreater than 2.311 and between about 6 and about 14 mol per cent carbondioxide, the carbon dioxide. concentration of the total feed being morethan inthe fresh feed.

References Cited in the tile of this patent UNITED STATES PATENTS2,251,554 Sabel et al Aug. 5, 1941 2,360,787 Murphree et al Oct. 17,1944 2,486,895 Watson Nov. 1, 1949 2,614,114 Krebs Oct. 14, 19522,631,159 Keith Mar. 10, 1953

